Process for treatment of coke oven light oil

ABSTRACT

A CATALYTIC PROCESS IS PROVIDED FOR EFFECTING SELECTIVE HYDROGENATION AND HYDROCRACKING OF COKE OVEN LIGHT OIL CONTAINING 10 TO 50% PRIMARY OIL. HEATING THE LIGHT OIL WITH HOT RECYCLE AND THE HYDROGEN, AND IMMEDIATE INTRODUCTION INTO CONTACT WITH HYDROGENATION CATALYST AT UPFLOW LIQUID PHASE CONDITIONS SUBSTANTIALLY REDUCES THE COKE-FORMING TENDENCIES OF NONAROMATIC UNSATURATED COMPONENTS IN THE LIGHT OIL. SUBSEQUENT CATALYTIC HYDDRODEALKYLATION PRODUCES IMPROVED YIELDS OF HIGH PURITY AROMATICS. IN ONE EMBODIMENT PROVISION IS MADE FOR THE RECOVERY OF NAPHTHALENE.

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' 7 ATTORNEYS United AStates Patent Office 3,564,067 PROCESS FORTREATMENT OF COKE OVEN LIGHT OIL Walter Brenner, Wayne, and Louis C.Doelp, Jr., Glen Mills, Pa., assignors to Air Products and Chemicals,Inc., Philadelphia, Pa., a corporation of Delaware Filed May 6, 1968,Ser. No. 726,883

Int. Cl. Cg 9/16, 7/00; C07c 3/58 U.S. Cl. 260-672 10 Claims ABSTRACT OFTHE DISCLOSURE BACKGROUND OF THE INVENTION The present invention isconcerned with purification and conversion of coke oven light oils toobtain therefrom aromatic hydrocarbons of enhanced value and insubstantially pure form.

The light oils produced as by-products in coke oven plants containchiefly benzene and other benzene-type hydrocarbons in relatively highproportions; however, these light oils are contaminated with a widevariety of hydrocarbonaceous materials in small amounts includingparaflins, olefins and naphthenes as well as significant quantities ofsulfur-containing compounds. The aromatic hydrocarbons being of highercommercial value are preferably recovering in high yield and in highpurity in order to realize the full value thereof. Commercial processesare in successful operation for the separation and recovery of purifiedaromatic portions from so-called secondary light oils.

In conventional practice the light oil separated from coke-oven gas issubjected to rectification to separate out an overhead fractioncomprising the secondary light oil and there is also recovered a bottomsfraction of so-called primary light oil, sometimes referred to as middleoil. This higher-boiling fraction contains a portion of the xylenes,naphthalene With or Without some methyl naphthalenes, and varioushydrocarbons boiling largely in the range of from xylene to about 250 C.including thiophenes, styrenes and other unsaturates together wtihindene, coumarones and other easily polymerizable gumforming materials.Because of the coking tendencies of the naphthalenes and other bicyclicspresent in the promary light oil as well as of the gum-formingunsaturated hydrocarbons, hydrogenative treatment of these primary lightoils over catalyst for recovery of purified aromatic hydrocarbonspresents operational difficulties. For these reasons, recovery ofaromatic hydrocarbons from coal tar oils, to the extent that is beingpracticed, is largely carried out by chemical treatment or to a lesserdegree by thermal (non-catalytic) hydrotreating methods under conditionsresulting in comparatively low yields of the desired aromatichydrocarbons due to extensive cracking of hydrocarbons in the charge togases and other low molecular weight compounds. Only incomplete removalof sulfur compounds is achieved by these methods.

3,564,067 Patented Feb. 16, 1971 SUMMARY OF THE INVENTION An object ofthe present invention, accordingly, is to provide an improved processfor handling of the primary light oil components to obtain valuablearomatic hydrocarbons therefrom in economically attractive yields.

In accordance with the invention coke oven light oils containing atleast 10% by volume of primary light oil including naphthalene and otherbicyclic compounds as well as easily polymerizable gum-forming materialsare pretreated with hydrogen over hydrogenative catalyst at processingconditions effecting selective hydrogenation and selective hydrocrackingwhereby subsequent processing to high yields of aromatics is effectedwith substantially reduced coke formation. Particularly beneficialresults are obtained with charge stocks containing l0 to 50% primarylight oil and 90 to 50% secondary light oil. Under the selectedconditions some of the unsaturates, particularly readily polymerizablecomponents, are hydrogenated. In one embodiment, the process conditionsare controlled, particularly as to temperature and hydrogen partialpressure, so as to favor saturation of only one of the rings of thebicyclic components present in the feed. The hydrogenated eluent isdistilled to remove (I) as bottoms, a higher boiling fraction containingsubstantially all unconverted naphthalene, and (II) a lower boilingoverhead fraction, including the monocyclic hydrocarbons. At least aportion of the higher boiling fraction (I) is reheated and recycled tothe bottom of the hydrogenation reactor, thus supplying heat to theentering fresh light oil charge and bringing the total charge (freshplus recycle) to desired reaction temperature. The monocyclic fraction(Il) is treated to effect dealkylation of alkyl aromatics,hydrodesulferization and selective catalytic hydrocracking for recoveryof enhanced amounts of extremely pure benzene with or without separaterecovery of toluene and/or Xylenes.

In another embodiment of the present invention, process conditions aresuch as to favor retention of naphthalene such that the higher boilingfraction (I) is fractionated to obtain a naphthalene heart cut. Theremaining portion of fraction (I) is then reheated and recycled to thebottom of the hydrogenation reactor.

An important aspect of the present invention lies in the novelarrangement for handling of the fresh feed containing unsaturates whichreadily polymerize to gums and tars on heating or prolonged storage. Toavoid polymer deposition in the lines and on hot Contact surfaces thehydrogenative pretreatment is effected at conditions such that the freshfeed is in substantially liquid phase and moves in upward -flow throughthe bed of hydrogenation catalyst, the charge being brought from ambienttemperature to desired inlet temperature by contact with preheatedrecycle bottoms stream (I) from fractionation of the hydrogenatedefliuent, such recycle stream being in a ratio in excess of one volumeper volume of fresh feed. The maintenance of the normally liquid portionof the charge in substantially liquid phase coupled with upow operationappear as essential features for successful extended performance.Nominally similar amounts of reactants at similar conditions provide asimilar degree of liquid phase in downow operation; however, the percentliquid occupancy at the described conditions is effectively higher atupflow conditions. It is desirable, therefore, to select conditionsrelated to the specific charge stock which will provide the lowestpractical reaction temperature, the highest practical pressureconsistent with the desired result and sound economics, and the lowesteffective quantity of added hydrogen.

The quantity of hydrogen supplied to the pretreating reactor should bein excess of the stoichiometric requirement for saturating oleins andoleiinic side chains in the feed. In preferred practice, the amount ofhydrogen supplied is less than one mole of hydrogen per mole of freshfeed and is generally in the order of about 0.5 mole of hydrogen permole of fresh feed for a charge comprising to 50 percent by volume ofprimary light oil. When benzene is the main desired product, theoperating conditions are relatively severe, involving temperatures aboveabout 285 C. (285-370 C.) where hydrogenation of monocyclic aromaticstends to occur, and the added hydrogen should be below one mole per moleof fresh feed. When less severe operating conditions are employed, suchas in the embodiment concerned with the separation and recovery ofsignificant quantities of naphthalene and particularly when higherconcentrations of primary light oil are present in the fresh feed, e.g.,with hydrogenation temperatures between 250 C. and 285 C., it ispossible to add as much as up to 2 moles of hydrogen per mole of freshfeed. In this latter instance the driving force to completehydrogenation may be less; but larger quantities of hydrogen-consumingcomponents, e.g., up to 40% indenes, naturally can account for largeramounts of added hydrogen.

BRIEF DESCRIPTION OF THE DRAWINGS The operation of the invention andcertain of the advantages thereof will be understood and appreciatedfrom the description which follows read in connection with theaccompanying drawings in which:

FIG. 1 is a `flow diagram illustrating the treatment of coke oven lightoil for the recovery of puried monocyclic aromatic hydrocarbons: and

FIG. 2 is a flow diagram of an embodiment illustrating the treatment ofcoke oven light oil wherein naphthalene is recovered in addition toother purified aromatic hydrocarbons.

DESCRIPTION OF THE PREFERRED EMBODIMENTS Referring now to FIG. l, thefresh feed of coke oven light oil at ambient temperature is introducedthrough conduit 10 leading to the bottom of hydrogenation reactor 11.This fresh feed includes the usual secondary light oil and up to about50% by volume of higher boiling components, including naphthalenes, inthe primary light oil boiling range. A heavier hot oil containing addedhydrogen is introduced through conduit 12-the source of this heavier oilwill be hereinafter explained-for intimate admixture with the fresh feedin conduit 10; and the admixture of fresh feed, heavier hot oil andhydrogen is introduced into the bottom of reactor 11. Thus, dilution ofthe fresh charge, establishment of the proper reaction temperature andcommencement of the reaction occur substantially simultaneously.

Reactor 11 contains sulfided cobalt molybdate catalyst supported onalumina. The preferred catalyst is one containing, prior to sulfidation,10 to 20%' by weight of the oxides of cobalt and molybdenum; the M003being from 3 to 5 times that of the COO by weight. Suldation of thecatalyst can be effected by pretreatment with H28. However, commonpractice with the sulfur-bearing stocks employed in this inventioninvolves simply allowing the sulfur in the charge to effect thesulfidation of the catalyst. A typical cobalt molybdate catalyst for usein the hydrogenation reactor is that described in Example I of U.S. Pat.No. 3,207,802. Catalysts containing small amounts of other metal oxidesor sulfides of the iron group, particularly nickel, in addition tocobalt may also be employed, such as catalysts of the type described inU.S. Pat. No. 2,880,171.

In reactor 11, under the selected process conditions, hydrogenation ofpolymerizable unsaturates, such as styrene, indene, anddicyclopentadienes, is effected. Saturation of one of the rings of thenaphthalenes present to form tetrahydronaphthalene-type compounds mayalso occur. In addition, 50 to 80% of the ring sulfur compounds(thiophenes for example) are hydrogenated, facilitating subsequentdesulfurization. A portion of these sulfur compounds may also behydrocracked in this operation. The conditions employed in reactor 11include temperatures in the range of about Z50-370 C. and pressures ofabout 50 to 100 atmospheres gauge. Since the desired selectivehydrogenation requires only a short residence time, high oil throughputrates can be employed including fresh feed volume space rates (LHSV) inthe order of 0.5 to 5.0. Preferred conditions are oil inlet temperaturesof about 300 C. and pressure of 70 atmospheres at an hourly oil spacerate (fresh feed only) of about l volume per volume of catalyst in thereactor. A volume space rate in the range of 0.8 to 20 yLHSV for thetotal oil feed is consistent with the described operation.

Under the conditions employed, a major portion of the oil will be inliquid phase during hydrogenation in the reactor 11 and the still liquidproduct as well as gases and vaporized products will be discharged atthe top of the reactor through conduit 13. The mixed-phase effluent,preferably after cooling to about 60 to 90 C. as indicated at 14 is sentto liquid-gas separator or flash drum 15. From the separator 15 there iswithdrawn through conduit 16 a vapor overhead composed of hydrogen andlight hydrocarbon gases up to about C5. The high pressure liquid fromhigh pressure liquid-gas separator 15 passes through conduit 17 andpressure reducing valve 18 into a low pressure liquid-gas separator 19from which low pressure gas is vented through conduit 20 and the gasfreeliquid passes through conduit 21 to fractionator 22. Separator 19 isoperated at any appropriate pressure of less than about 7 atmospheresgauge and preferably in the range of about atmospheric to about 3atmospheres pressure. Suitable valving, liquid level controllers andrelated instrumentalities are employed as required.

IFractionator 22 is operated under conditions and at a. cut point tosparate out a liquid fraction, including naphthalenes, from the vaporoverhead which includes monocyclic aromatcs. If the fractionator 22 isoperated at substantially atmospheric pressure, this cut point will beat a level in the range of about C. to 120 C., so that at least themajor portion of th indane and tetralin will be in the liquid fraction.

The vapor overhead from fractionator 22 is withdrawn through conduit 23and passes through condenser 24 in which essentially all of the C5+components are condensed to liquid form. The effluent from condenser 24is separated in a liquid-gas separator 25 with uncondensed vaporsvented, or otherwise disposed of, through conduit 26. A portion of thecondensate can be returned from the separator through conduit 27 asliquid reflux to fractionator 22 while the remaining portion of thecondensate is passed from the separator through conduit 28 topressurizing pump 29. The pressurized liquid from pump 29 is then passedthrough conduit 30 and combined with the vapor overhead from separator15 in conduit 16 for further processing, as will be further describedbelow.

The liquid fraction from fractionator 22 is withdrawn through conduit 31and passes through pump 32 into conduit 33. Any excess liquid fractionbeyond the requirements for recycle use may be withdrawn from the system`as from line 31 by way of valved outlet 50. To assist in the operationof fractionator 22, a reboiler may be provided at the bottom thereof.Hydrogen-containing gas is added via conduit 34 to the liquid fractionin conduit 33, and the mixture of oil and gas is then heated to arequired temperature at 3S before being recycled through conduit 12 tothe bottom of the pretreat reactor 11. Only a small amount of hydrogenneed be added to the recycle oil stream, i.e., sufficient hydrogen toeffect the saturation of polymerizable olefins and no more than one ofthe rings of naphthalene compounds contained in the fresh light oil feedadmitted through conduit 10. As indicated above, as little as about 0.5mole of hydrogen per mole of fresh oil charged is normally adequate.However, larger amounts of hydrogen can be employed, up to that amountwhich would tend to favor undesired hydrogenation of monocyclicaromatics and/or interfere with the substantially liquid phase operationunder the conditions employed in reactor 11.

The' recycled oil admitted to reactor 11 provides the principal sourceof the heat required for initiation of the hydrogenation reaction inreactor 11. Accordingly, the stream in conduit 12 is brought to atemperature sul'licient to provide in the mixture thereof with the freshfeed from conduit the desired inlet temperature at the bottom of reactor11. The fresh feed charged through conduit 10 contains unsaturatedhydrocarbons which tend to polymerize at elevated temperature (e.g., atabout 150 C.) and cause fouling by deposit of tarry and otherpolymerized substances on catalyst and other surfaces contacted thereby.This problem of fouling equipment and catalyst is overcome by heatingthe fresh feed directly and only with the hot stream from conduit 12which is admixed therewith. Since the hot oil in conduit 12 has alreadyIbeen subjected to hydrogenation it is relatively free of troublesomeunsaturates and provides an excellent diluent and wash oil for the freshfeed. Relatively long catalyst life and substantial freedom from foulingof equipment are obtained when the recycled hot oil in conduit 12 isemployed at a ratio of at least 1 and preferably at about 3 volumes pervolume of fresh feed. Higher recycle oil ratios of course can beemployed but have no particular addedadvantage and tend unnecessarily toreduce the useful capacity of the reactor and other equipment. The useofhot refractory oils in another relation for supplying heat directly toan oil charge containing polymerizable unsaturates is described in U.S.Letters Patent No. 3,216,924.

The vapor overhead from separator 15 passes through conduit 16 and iscombined with the net overhead from fractionator 22 which passes throughconduit 30. The combined material, free of polymerizable contaminants,passes through conduit 36 to the hydrodealkylation system showndiagrammatically at 37. Eiiiuent from the hydrodealkylation system 37 iswithdrawn through line 49 for separation and recovery ofhydrogen-containing gas and product aromatics by means not shown. Thehydrogen-containing gas necessary to the hydrodealkylation reaction isintroduced through conduit 38 and cornbined with the material in conduit36. Typically, this hydrogen-containing gas comprises high pressurerecycle gas, from the hydrodealkylation system 37, supplemented by suchmake-up hydrogen gas as may be desired or required. For example, thehydrogen-containing gas in conduit 38 should contain at least 70% freehydrogen in an amount in the range of 4 moles to 10 moles of hydrogen,usually about 6 moles, per mole of aromatic hydrocarbons in conduit 36.

In practice, the hydrodealkylation methods described in U.S. LettersPatent No. 3,081,259 may be utilized at 37. As therein described, thecharge is introduced into at least one reactor together with hydrogen atan inlet temperature of about S90-630 C. and contacted at 30 to 70atmospheres with high activity chromia-alumna catalyst so that duringreaction 'temperatures of at least 630 C. are reached. Space rates arepreferably employed such that under the reaction conditions the nominalresidencetime of the material is less than three minutes. At theseconditions sulfur-containing compounds are substantially completelyconverted as by hydrocracking, so that any sulfur is in the form ofhydrogen sulfide; non-aromatic compounds are hydrocracked to lighthydrocarbons; and a substantial portion of alkyl aromatics ishydrodealkylated to benzene. The effluent from the hydrodealkylationstep is normally flashed to remove HZS, H2 and the lighter hydrocarbons.Such gaseous portion may be freed of HZS and then utilized with orwithout other purication as at least a portion of thehydrogen-containing gas 1ntroduced with the charge of hydrodealkylation.The nongaseous portion of the flashed-effluent is distilled andotherwise treated to effect separation and recovery of purifiedmononuclear aromatics, chiefly benzene.

The embodiment shown in FIG. 2 for the recovery of naphthalenc, as wellas other puried aromatic hydrocarbons, is similar to that illustrated inFIG. 1. For the sake of simplicity, items which are identicalstructurally in FIGS. 1 and 2 have been designated with the samenumeral. With the following exceptions the process steps and conditionsemployed in the embodiment illustrated by FIG. 2 are identical withthose for the embodiment illustrated in FIG. 1.

Since naphthalene is recovered in the embodiment shown in FIG. 2, lesssevere process conditions are utilized in hydrogenation reactor 11 thanin connection with the embodiment shown in FIG. 1. The desired selectivehydrogenation of polymerizable unsaturates, without extensive saturationof naphthalene rings, can be effected by operating reactor 11 at atemperature in the range of 250285 C. and at a pressure of from about 50to 100 atmospheres. Preferred conditions are oil inlet temperatures ofabout 260 C. and a pressure of 70 atmospheres at an hourly oil spacerate (total oil including recycle) of 3 volumes per volume of catalystin the reactor.

The actual recovery of naphthalene is effected by passing the liquidfraction withdrawn from fractionator 22 through conduit 39 to anotherfractionator 40. A naphthalene heart cut, e.g., boiling between 217-219C., is recovered from fractionator 40 in conduit 41 and furtherprocessed, by means such as a stripper tower, not shown, in a mannersuitable for the separation and recovery of pure naphthalene. Theoverhead fraction from fractionator 40 is passed through line 42 tooverhead condenser 43. Uncondensed vapor may be vented through line 44and condensate removed through line 45 from which a suitable recyclestream passes through line 46 for re-entry to fractionator 40. Theremaining portion of the condensate from line 45 passes through line 47for combination with the 'bottoms fraction, from fractionator 40, passedthrough line 48. The combined streams from lines `47 and 48 are passedthrough line 31 for return as recycle to reactor 11. Valved outlet 50from line 31 is provided to permit withdrawal of any liquid in excess ofthat required for recycle purposes.

A fuller understanding of the invention will be had from the followingexamples depicting the application thereof in several preferredembodiments. It is to be understood that these examples are forillustrative purposes only and are not intended as limiting.

lExample I Fresh coke oven light oil feed containing 20% by weightprimary light oil and by weight secondary light oil was introduced atambient temperature into the bottom of a hydrogenation reactor. Thecomposition of this fresh feed is shown in Table l.

A heavier hot oil having the composition shown in Table 2 was also addedin an amount of 3.3 volumes per volume of fresh coke oven light oil feedto the bottom of the hydrogenation reactor together with 0.55 mole ofhydrogen per mole of fresh feed.

TABLE 2 Material: Wt. percent C11C12+ aromatics 91.23 Naphthalene 7.37Indanes 0.50 Indenes 0.44 Benzene 0.30 Toluene 0.14 Xylenes 0.02

The fresh feed together with the added heavier hot oil and hydrogen werecontacted with alumina-supported presulded cobalt molybdate catalyst,having a composition of about 82% by weight alumina, 3% by weight COO,15% by weight M003 and a bulk density of 0.7, at an average temperatureof 313 C., a pressure of 69 atmosphere absolute and a fresh feed liquidhourly space velocity of 0.9.

The mixed-phase effluent from the hydrogenation reactor was then sent toa liquid-gas separator. The gaseous overhead fraction was passeddirectly to a chromia-alumina catalyst reactor for hydrodealkylation asindicated below.

The liquid fraction from the liquid-gas separator was reduced inpressure to 2 atmospheres absolute and sent to a fractionating column,containing theoretical plates. The composition of the condensed overheadfrom the fractionating column is shown in Table 3.

TABLE 3 Material: Wt. percent Benzene 62.13 Toluene 15.02 Indanes 9.81Xylenes 4.16 C9 to C11 aromatics 3.17 Ethyl benzene 1.95 Naphthalenes1.33 C2-C6 parafns 1.17 Tetralin 1.04 Diphenyl 0.22

The liquid fraction from the fractionating column, having a compositionshown in Table 2, was recycled and added to the fresh feed introduced tothe hydrogenation reactor.

The condensed overhead from the fractionating column together with thegas from the liquid-gas separator were l combined with fresh hydrogen ina ratio of 1.4 moles of hydrogen per mole of aromatics, and 7.8 moles ofrecycle gas, containing a minimum of 70 mole percent hydrogen, (from thehydrodealkylation reactor) per mole of aromatics. The combined mixturewas then contacted with a chromia on alumina catalyst at a temperatureof 635 C. for a residence time of 36.9 seconds. The chromia on aluminacatalyst was a commercially available catalyst containing a nominal 20%by weight Cr2O3. Similar catalysts having 15 to 25% by weight of Cr203impregnated on alumina which has a surface area of 100 to 200 squaremeters per gram before impregnation may be ernployed with generallysimilar results.

Analysis of the resulting liquid product from the hydrodealkylationshows 91.50 wt. percent benzene, 6.37 wt. percent toluene and minoramounts of the materials shown in Table 4.

TABLE 4 Material: Wt. percent Diphenyl 1.05 Ethyl benzene 0.59Naphthalene 0.24 Xylenes 0.12 C2 to C5 parains 0.10

C9 to C10 aromatics 0.03

Benzene was separated from non-benzene portions of the liquid product byfractionation to obtain 99.9 mole percent benzene.

As seen by this example very little naphthalene is in the vapor overheadobtained from the fractionating column and practically no benzene isrecycled.

Example II Fresh coke oven light oil feed containing approximately 50%by weight primary light oil and 50% by weight secondary light oilconstituted the fresh feed portion of the charge to the hydrogenationreactor. This portion of the charge, boiling in the range of 60.5 C. to260 C., had the composition shown in Table 5.

TABLE 5 Material: Wt. percent Benzene 38.39

Toluene 7.79 Xylenes 2.81 Ethyl benzene 0.07 Styrene 1.14 C9+ aromatics8.34 Coumarone 1.50

Indanes 0.37

Indenes 16.28

Tetralin 1.04 Naphthalene 19.96 Methyl naphthalenes 1.19 Benzothiophenes0.79 Thiophene 0.29 Other 0.04

To this fresh feed portion of the charge there is added, in a ratio of1.2 to 1 volume of fresh feed, a stream of recycle oil, from thenaphthalene separator, to which is added fresh hydrogen in an amountequivalent to 1.65 moles per mole of fresh feed. The recycle oil has thecomposition shown in Table 6.

TABLE 6 Material: Wt. percent C6C8 aromatics 2.07 09+ aromatics 20.43Indanes 67.52

Tetralin 4.99 Methyl naphthalene 4.42 Benzothiophenes 0.57

The fresh feed portion of the charge was introduced at ambienttemperature and raised to the reactor inlet temperature of about 260 C.by admixture with the heated recycle oil-hydrogen stream. The admixedstream was directly introduced to the upllow reactor into contact withan alumina-supported sulded cobalt molybdate catalyst similar to thatdescribed in Example I.

The reaction conditions included an operating pressure of 69 atmospheresabsolute, a liquid hourly space velocity of 0.9, based on fresh feed,and the relatively low average temperature of about 264 C.

The effluent from the hydrogenation reactor was separated in a flashingoperation into a vapor overhead, forming part of the charge to asubsequent hydrodealkylation stage or other use, and a liquid bottoms.The liquid portion amounted to 98.25% by weight of the total charge tothe hydrogenation reactor. After pressure reduction to about 33 p.s.i.a.this liquid portion was distilled in a fractionation column containing25 theoretical plates into an overhead fraction amounting to 39.11% byweight of total charge suitable as charge to a subsequenthydrodealkylation stage, and a bottoms fraction amounting to 59.14% byweight of the total charge to the hydrogenation stage.

The bottoms fraction was distilled in a second fractionation in a higheiciency fractionation column. An overhead fraction, comprising mainlyC7-C9 aromatics, and a bottoms fraction were removed and recombined to-form the above-mentioned recycle oil used as a portion of the charge tothe hydrogenation reaction. There was also separated and recovered anintermediate fraction containing substantially only the naphthalene inan amount equivalent to 12.54 Weight percent of the fresh feed to thehydrogenation reactor.

- The condensed overhead from the first fractionator is combined withthe gas from the high pressure flash, hydrogen is added and the streamis then processed in the hydrodealkylation operation as described inExample I. Benzene, toluene and xylene ,(BTX) are recovered as lowsulfur product having purity better than 99.9 mole percent.

In accordance with operation of this example desired aromatics arerecovered in amounts, based on the fresh feed, of 12.54 :weight percentnaphthalene and 76.96 percent by weight as BTX.

Thus, the present invention provides a process for the purification ofcoke oven light oil contaminated with nonaromatic hydrocarbons. Thepresent invention is not only able to handle charge stocks which cannotbe processed by other catalytic systems for any practical period of onstream time, but is able to obtain an aromatics recovery which is far inexcess of that obtained from similar stocks treated by either chemicalor thermal means. The recovery of low sulfur containing products isanother unique feature of the present invention. This process utilizespartial product recycle as diluent for, and the preheating of, freshcoke oven light oil immediately prior to subjecting the charge stock tohydrogenation in an upflow reactor operated under substantially liquidphase conditions. Selective fractionation in the process prevents anysignificant amount of bicyclic aromatics from reaching the finalhydrodealkylation system. Selective hydrogenation and hydrocracking ofnaphthalene and other bicyclics is achieved in one embodiment of theprocess. In another embodiment, naphthalene is recovered as a product byemploying somewhat milder operating conditions for hydrogenating anyunsaturates with substantially less hydrogenation and hydrocracking ofthe naphthalene and using a second fractionator to obtain a naphthaleneheart cut.

What is claimed is:

1. The process for the recovery of aromatic hydrocarbons from coke ovenlight oil containing at least by volume of primary light oil, whichcomprises:

(a) admixing the coke oven light oil at ambient temperature and a hotstream comprising a recycle liquid fraction having an initial boilingpoint in the range of 100-120 C. in an amount in the range of 1 to 3volumes per volume of light oil and hydrogen in an amount in the rangeof 0.5 to 2.0 mols of hydrogen per mole of light oil and obtaining anadmixture having a temperature in the range of from about 250-300 C.

(b) introducing said admixture under substantially liquid phaseconditions into an upilow hydrogenation reactor into contact rwithhydrogenation catalyst and effecting hydrogenation of readilypolymerizable components,

(c) subjecting a mixed phase eiuent from the hydrogenation reactor toliquid-gas separation into a vapor overhead fraction and a liquidfraction,

(d) fractionating the liquid fraction into a lower boiling fractionhaving an end boiling point in the range of about 100-120 C. and ahigher boiling liquid fraction,

`(e) recycling at least a portion of the higher boiling liquid fractionas said recycle liquid fraction,

(f) admixing at least a portion of the vapor overhead fraction from (c),at least a major portion of the lower boiling fraction from (d), andhydrogen to form an hydrodealkylation charge stream,

|( g) subjecting the hydrodealkylation charge stream tohydrodealkylation at hydrodealkylation conditions 10 including contactwith chromia-alumina catalyst, and

(h) recovering substantially pure aromatic hydrocar- =bons from theeffluent from hydrodealkylation.

2. The process of claim 1 in which the coke oven light oil contains l0to 50 percent by Volume of primary light oil.

3. The process of claim 1 in which the coke oven light oil, recycleliquid fraction and hydrogen-containing gas admixture is contacted withcobalt molybdate catalyst in the hydrogenation reactor at a temperaturein the range of about 250 to about 370 C., a pressure of 50 toatmospheres gauge and a space rate of between 0.5 and 5.() LHSV, basedon the coke oven light oil.

4. The process of claim 3 in which the cobalt molybdate catalyst is asulfided cobalt molybdate catalyst supported on alumina containing,prior to sultidation, 10 to 20% by weight of the oxides of cobalt andmolybdenum and in which M003 is present in an amount 3 to 5 times thatof the COO by weight.

5. The process of claim 1 in which the mixed phase eluent from thehydrogenation reactor is cooled to a temperature between 60 and 90 C.,before being subjected to liquid-gas separation.

6. The process of claim 1 in which the fractionation in (d) is effectedat a pressure in the range of 1 to 3 atmospheres and with a cut point inthe range of about 100 to about 120 C. at atmospheric pressure.

7. The process of claim 1, in which the hydrodealkylation conditionsinclude a temperature in the range of 590 to 630 C. and a pressure inthe range of 30 to 70 atmospheres gauge with resulting dealkylation ofalkyl aromatics to aromatic hydrocarbons; and fractionating the effluentfrom hydrodealkylation to separate and recover the aromatichydrocarbons, including benzene.

8. The process of claim 1 in which naphthalene is recovered as productby taking a heart cut of a fraction boiling in the range of between 217and 219 C. from a second fractionation of the higher boiling fraction in(d).

9. The process for the recovery of high purity aromatic hydrocarbons,including; benzene and naphthalene, from coke oven light oil containing10 to 50 percent by volume of primary light oil, which comprises: (a)admixing the coke oven light oil having an ambient temperature with ahot stream of recycle liquid and hydrogen-containing gas, said recycleliquid being employed in an amount in the range of 1-3 volumes pervolume of coke oven light oil and the hydrogen in saidhydrogen-containing gas being employed in an amount in the range of 0.5to 2.0 moles of hydrogen per mole of coke oven light oil, said admixturehaving a temperature in the range of 250 to 285 C. obtained with saidhot stream being the direct and only source of heating for the coke ovenlight oil;

(b) introducing the admixture into an upow hydrogenation reactor intocontact with hydrogenation catalyst at a temperature between 250 and 285C., a pressure between 50 and 100 atmospheres gauge and a space ratebased on the coke oven light oil of between 0.5 and 5.0 LHSV, saidcontacting being effected with at least said coke oven light oil beingin substantially liquid phase;

(c) subjecting the mixed phase effluent from the hydrogenation reactorto liquid-gas separation to obtain a vapor overhead fraction and aliquid fraction;

(d) fractionating the liquid fraction into a lower boiling fraction anda higher boiling fraction, said fractionating having a cut pointequivalent to between about 100 and about 120 C. at atmosphericpressure;

(e) subjecting the higher boiling fraction to further fractionationtoseparate out and recover a heart cut boiling in the range of between 217and 219 C. and comprising mostly naphthalene;

(f) recombining the fractions, minus said heart cut,

11 from said further fractionation and forming a recycle liquid;

(g) admixing said recycle liquid and hydrogen-containing gas and heatingsaid last mentioned admixture to provide said hot stream in (a);

(h) combining said vapor overhead fraction of (c) and said lower boilingfraction of (d) with hydrogen to form a hydrodealkylation charge stream;

(i) contacting the hydrodealkylation charge stream with a dealkylationcatalyst in a hydrodealkylation 1 reactor at hydrodealkylationconditions including a temperature between about 590 and 630 C. and apressure between 30 and 70 atmospheres gauge; and (j) recoveringaromatics, including high purity benzene, from the hydrodealkylationreactor eflluent. 10. The process of claim 1 in which said dealkylationcatalyst is chromia on alumina and benzene is recovered from the reactoreffluent by fractionation.

References Cited UNITED STATES PATENTS Donovan et al. 208-216 Butler etal. 208-255 Butler et al. 208-89X Maerker et al 260-674 Watkins 20S-143McKinney et al. 208-143 Assistant Examiner U.S. Cl. X.R.

P34050 UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No.3,56#,O67 Dated February 16, 1971 lnventods) Walter Brenner and Louis CDoelp, Jr.

It is certified that erro: appears in the above-identified" patent andthat said Letters Patent are hereby corrected as shown below:

r- Column l, line bf2, "recovering" should read --recovered Column 1,line 59, "promary" should read --primary Column 2, line 314,"hydrodesulferization" should read --hydrodesulfurization Column il,line 37, l'sparate" should read separate column u, line u2, "th" shouldread the-- Column 5, line T5, "of" should read to Column 7, line 19,"atmosphere:l should read atmospheres-- Column lO, line lll, deletesemi-colon Column lO, line 68, "fraotonating" should read --fractionaColumn l1, line 16, "claim l" should read --claim 9 Signed and sealedthis 18th day of April 1972.

(SEAL) Attest:

EDWARDFLFLETCHERJR. ROBERT GOTTSCHALK At testing Officer Commissioner ofPatents

